U.S. patent application number 17/497071 was filed with the patent office on 2022-03-31 for process for the preparation 1,1,1,2,2-pentafluoropropane.
This patent application is currently assigned to Mexichem Fluor S.A. de C.V.. The applicant listed for this patent is Mexichem Fluor S.A. de C.V.. Invention is credited to Stephen Andrew Flaherty, Clive Robert Giddis, Sheryl Louise Johnson.
Application Number | 20220098131 17/497071 |
Document ID | / |
Family ID | 1000006015271 |
Filed Date | 2022-03-31 |
United States Patent
Application |
20220098131 |
Kind Code |
A1 |
Johnson; Sheryl Louise ; et
al. |
March 31, 2022 |
PROCESS FOR THE PREPARATION 1,1,1,2,2-PENTAFLUOROPROPANE
Abstract
The present invention provides a process for preparing
1,1,1,2,2-pentafluoropropane (245cb), the process comprising gas
phase catalytic dehydrochlorination of a composition comprising
1,1,1-trifluoro-2,3-dichloropropane (243db) to produce an
intermediate composition comprising
3,3,3-trifluoro-2-chloro-prop-1-ene (CF.sub.3CCI.dbd.CH.sub.2,
1233xf), hydrogen chloride (HCl) and, optionally, air; and gas
phase catalytic fluorination with hydrogen fluoride (HF) of the
intermediate composition to produce a reactor product composition
comprising 245cb, HF, HCl and air; wherein the process is carried
out with a co-feed of air.
Inventors: |
Johnson; Sheryl Louise;
(Cheshire, GB) ; Flaherty; Stephen Andrew;
(Cheshire, GB) ; Giddis; Clive Robert; (Cheshire,
GB) |
|
Applicant: |
Name |
City |
State |
Country |
Type |
Mexichem Fluor S.A. de C.V. |
San Luis Potosi |
|
MX |
|
|
Assignee: |
Mexichem Fluor S.A. de C.V.
San Luis Potosi
MX
|
Family ID: |
1000006015271 |
Appl. No.: |
17/497071 |
Filed: |
October 8, 2021 |
Related U.S. Patent Documents
|
|
|
|
|
|
Application
Number |
Filing Date |
Patent Number |
|
|
16855500 |
Apr 22, 2020 |
11155506 |
|
|
17497071 |
|
|
|
|
15745529 |
Jan 17, 2018 |
10669219 |
|
|
PCT/GB2016/052133 |
Jul 14, 2016 |
|
|
|
16855500 |
|
|
|
|
Current U.S.
Class: |
1/1 |
Current CPC
Class: |
C07C 17/206 20130101;
B01J 23/06 20130101; C07C 17/25 20130101; C07C 19/08 20130101; C07C
21/18 20130101; B01J 23/26 20130101 |
International
Class: |
C07C 17/20 20060101
C07C017/20; B01J 23/06 20060101 B01J023/06; B01J 23/26 20060101
B01J023/26; C07C 17/25 20060101 C07C017/25 |
Foreign Application Data
Date |
Code |
Application Number |
Jul 17, 2015 |
GB |
1512557.8 |
Claims
1. A process for preparing 1,1,1,2,2-pentafluoropropane (245cb) and
2,3,3,3-tetrafluoropropene (1234yf), the process comprising: gas
phase catalytic dehydrochlorination of a composition comprising
1,1,1-trifluoro-2,3-dichloropropane (243db) to produce an
intermediate composition comprising
3,3,3-trifluoro-2-chloro-prop-1-ene (CF.sub.3CCl.dbd.CH.sub.2,
1233xf), hydrogen chloride (HCl), or 1233xf, HCl and air; and gas
phase catalytic fluorination with hydrogen fluoride (HF) of the
intermediate composition to produce a reactor product composition
comprising 245cb, 1234yf, HF, HCl and air; optionally wherein the
process is carried out with a co-feed of air.
2. The process according to claim 1 wherein the dehydrochlorination
step is carried out in a first reactor and the fluorination step is
carried out in a second reactor.
3-10. (canceled)
11. The process according to claim 2 wherein the intermediate
composition exits the first reactor and is fed directly to the
second reactor.
12. A process for preparing 1,1,1,2,2-pentafluoropropane (245cb)
and 2,3,3,3-tetrafluoropropene (1234yf), the process comprising:
gas phase catalytic dehydrochlorination in a first reactor of a
composition comprising 1,1,1-trifluoro-2,3-dichloropropane (243db)
to produce an intermediate composition comprising
3,3,3-trifluoro-2-chloro-prop-1-ene (CF.sub.3CCl.dbd.CH.sub.2,
1233xf) and hydrogen chloride (HCl); and gas phase catalytic
fluorination with hydrogen fluoride (HF) in a second reactor of the
intermediate composition to produce a reactor product composition
comprising 245cb, 1234yf, HF, HCl and air; optionally wherein the
process is carried out with a co-feed of air to the second
reactor.
13-18.
19. The process according to claim 12 wherein the intermediate
composition exits the first reactor and is fed directly to the
second reactor.
20. The process according to claim 12 wherein the catalytic
dehydrochlorination of 243db is carried out in the presence of HF
and the intermediate composition further contains HF.
21. The process according to claim 20 wherein the composition
comprising 243db additionally contains HF, with a molar ratio of
HF:243db of from 0.5:1 to 40:1, or from 1:1 to 15:1.
22. The process according to claim 20 wherein the contact time for
the composition comprising 243db and HF with the catalyst is from
0.5 to 200 seconds, or from 1 to 150 seconds, or from 1 to 100
seconds, or from 2 to 80 seconds, or from 8 to 60 seconds.
23. The process according to claim 2 wherein the molar ratio of
HF:1233xf in the second reactor is from 1:1 to 45:1, or from 2:1 to
20:1 or from 3:1 to 15:1, or wherein an additional feed of HF is
provided to the second reactor.
24. The process according to claim 2 wherein the contact time for
the composition comprising 1233xf, HCl and HF with the catalyst is
from 0.5 to 200 seconds, or from 1 to 150 seconds, or from 1 to 100
seconds, or from 2 to 80 seconds, or from 5 to 50 seconds.
25-26. (canceled)
27. The process according to claim 1 wherein the reactor product
composition is separated into a stream comprising 245cb, 1234yf and
HF and a stream comprising HCl and air.
28. The process according to claim 27 wherein the stream comprising
245cb, 1234yf and HF is separated into a 245cb-rich stream and a
HF-rich stream.
29. The process according to claim 28 wherein the 245cb-rich stream
contains 1234yf.
30. The process according to claim 28 wherein the HF-rich stream
contains 1233xf.
31. The process according to claim 27 wherein the separation is
achieved by distillation.
32. The process according to claim 27 wherein the 245cb-rich stream
is subjected to a scrubbing step in which residual HF is
substantially removed from the 245cb-rich stream to produce a
245cb-rich stream substantially free from HF.
33. The process according to claim 27 wherein the 245cb is
separated from the 1234yf to provide a substantially pure 245cb
product.
34. The process according to claim 33 wherein the separation is
achieved by distillation.
35. The process according to claim 1 wherein the catalytic
dehydrochlorination is carried out at a temperature of from 200 to
450.degree. C. and a pressure of from 0.1 to 30 bara, or at a
temperature of from 250 to 380.degree. C. and a pressure of from 1
to 20 bara, or at a temperature of from 300 to 350.degree. C. and a
pressure of from 5 to 20 bara.
36. The process according to claim 1 wherein the catalytic
dehydrochlorination is carried out in the presence of a bulk form
or supported catalyst comprising activated carbon, a zero-valent
metal, a metal oxide, a metal oxyhalide, a metal halide, or
mixtures thereof.
37. The process according to claim 36 wherein the metal is a
transition metal, an alkaline earth metal or aluminum.
38. The process according to claim 36 wherein the catalyst is based
on chromia, or a zinc/chromia catalyst.
39. The process according to claim 1 wherein the catalytic
fluorination is carried out at a temperature of from 200 to
450.degree. C. and a pressure of from 0.1 to 30 bara, or at a
temperature of from 250 to 420.degree. C. and a pressure of from 1
to 20 bara, or at a temperature of from 300 to 380.degree. C. and a
pressure of from 5 to 20 bara.
40. The process according to claim 1 wherein the catalytic
fluorination is carried out in the presence of a bulk form or
supported catalyst comprising activated carbon, a zero-valent
metal, a metal oxide, a metal oxyhalide, a metal halide, or
mixtures thereof.
41. The process according to claim 40 wherein the metal is a
transition metal, an alkaline earth metal or aluminum.
42. The process according to claim 40 wherein the catalyst is based
on chromia, or a zinc/chromia catalyst.
43. The process according to claim 1 wherein the HF in the reactor
product composition is at least partially recycled to the catalytic
dehydrochlorination of the composition comprising 243db and HF.
44. The process according to claim 28 wherein the HF in the HF-rich
stream is recycled to the catalytic dehydrochlorination of the
composition comprising 243db and HF.
45. The process according to claim 44 wherein the HF-rich stream is
separated into an HF stream and an organic stream, wherein the HF
stream is recycled to the catalytic dehydrochlorination of the
composition comprising 243db and HF.
Description
RELATED APPLICATIONS
[0001] This application is a continuation of co-pending application
Ser. No. 15/745,529 filed 17 Jan. 2018, which is a 371
international application Serial No. PCT/GB2016/052133 filed 14
Jul. 2016.
BACKGROUND OF THE INVENTION
[0002] The invention relates to a process for preparing
1,1,1,2,2-pentafluoropropane (HFC-245cb, referred to hereinafter as
245cb). In particular, the invention relates to a process for
preparing 245cb from 1,1,1-trifluoro-2,3-dichloropropane
(HCFC-243db, referred to hereinafter as 243db) via
3,3,3-trifluoro-2-chloro-prop-1-ene (HCFO-1233xf, referred to
hereinafter as 1233xf).
SUMMARY OF THE INVENTION
[0003] 245cb is a useful compound, not least as an intermediate in
the preparation of 2,3,3,3-tetrafluoropropene (HFO-1234yf, referred
to hereinafter as 1234yf). 245cb is mentioned as an intermediate in
the preparation of 1234yf in WO2009/125199. 245cb is also mentioned
in passing in other documents concerned with the preparation of
1234yf, such as WO2008/054781, WO2013/111911 and US2014/010750.
[0004] The listing or discussion of a prior-published document in
this specification should not necessarily be taken as
acknowledgment that the document is part of the state of the art or
is common general knowledge.
[0005] There is a need for an efficient and economic manufacturing
process for the preparation of 245cb. The subject invention address
this need by the provision of a process for preparing 245cb, the
process comprising gas phase catalytic dehydrochlorination of a
composition comprising 243db to produce an intermediate composition
comprising 1233xf, hydrogen chloride (HCl) and, optionally, air;
and gas phase catalytic fluorination of the intermediate
composition with hydrogen fluoride (HF) to produce a reactor
product composition comprising 245cb, HF, HCl and air; wherein the
process is carried out with a co-feed of air.
[0006] For the avoidance of doubt, the gas phase catalytic
dehydrochlorination comprises conversion of the 243db by
dehydrochlorination to 1233xf. Likewise, the gas phase catalytic
fluorination comprises conversion of 1233xf to 245cb by
fluorination. The conversion of 1233xf
(CF.sub.3CC.sub.1.dbd.CH.sub.2) to 245cb (CF.sub.3CF.sub.2CH.sub.3)
involves the addition of two fluorine substituents and one hydrogen
substituent to 1233xf. Put another way, this involves the addition
of HF and the replacement of the chlorine substituent with fluorine
substituent. Thus, the term fluorination (with HF) in the context
of the subject invention can be considered to include combined
fluorination and hydrofluorination reactions.
[0007] The above process may be carried out batch-wise or
continuously. Preferably, the process is carried out continuously.
The term "continuously" as use herein is intended to include
semi-continuous operation of the process wherein the process is
temporarily stopped, for example, to regenerate and/or replace the
catalyst used in the catalytic dehydrochlorination of 243db and/or
catalytic fluorination of 1233xf. Certain aspects of the invention
enable the cycle time between such catalyst regeneration and/or
replacement to be lengthened thereby improving the efficiency and
economy of the process.
[0008] The catalytic dehydrochlorination of 243db and the catalytic
fluorination of 1233xf of the process of the invention can be
carried out together in a single reactor.
[0009] In a preferred aspect, however, the catalytic
dehydrochlorination of 243db and the catalytic fluorination of
1233xf are carried out in separate first and second reactors,
respectively. Typically, there are advantages associated with the
use of separate reactors for these two reactions, including
modifying the conditions in each reactor to facilitate the
catalytic dehydrochlorination of 243db and the catalytic
fluorination of 1233xf, respectively. For example, a higher
pressure can be used for the catalytic fluorination of 1233xf
compared to the catalytic dehydrochlorination of 243db. Typically,
a somewhat higher temperature can be used in the second reactor
compared to the first reactor. One reason for such a difference in
reactor temperature is the preference for higher temperatures in
the second reactor to burn off any catalyst coking. This is
explained in more detail later in this specification. The use of
two reactors also helps different concentrations of HF and air to
be used in the catalytic dehydrochlorination and fluorination
reactions.
[0010] Whether a single or two reactors are used, any suitable
apparatus may be used. Typically, the apparatus is made from one or
more materials that are resistant to corrosion, e.g.
Hastelloy.RTM., Monel.RTM. or Inconel.
[0011] Regardless of whether one or two reactors is used, a key
feature of the invention is that it is carried out with a co-feed
of air. The inventors have surprisingly found that this prevents
and/or retards coking of the catalyst or catalysts used in the gas
phase catalytic dehydrochlorination of a 243db and/or the gas phase
catalytic fluorination of 1233xf (particularly the latter reaction)
without significantly impairing conversion and/or selectivity. Put
another way, the use of air has been found to significantly reduce
the rate of catalyst deactivation in the gas phase transformations
of the subject invention. This has the effect of lengthening cycle
time, which in turn has benefits of process efficiency and economy.
The use of an air co-feed is also believed to enable the process of
the invention to be conducted at higher temperatures with a given
catalyst. Without being bound by theory, the air co-feed is thought
to help burn coke at approximately the same rate at which it is
produced, thereby extending cycle time. This is believed to be
especially advantageous for the gas phase catalytic fluorination of
1233xf, which compound is fouling to the gas phase fluorination
catalyst and relatively difficult to convert to 245cb compared to
the typically more facile gas phase catalytic dehydrochlorination
of a 243db.
[0012] It is believed that it is the oxygen in the air that is
primarily responsible for the unexpected effects described in the
preceding paragraph. However, there are advantages to using air in
the process of the invention rather than oxygen or oxygen enriched
air. Use of air (e.g. atmospheric air) is both cheaper than using
oxygen or oxygen enriched air. It is also safer to handle air
compared to oxygen enriched air or, particularly, oxygen, due to
flammability issues. The concentration of oxygen in air (about 21
mol %) is also thought to be especially suitable for use in the
process of the invention, in terms of the combination of its
effectiveness to prevent and/or retard catalyst coking and its ease
of handling. For example, the air, in one embodiment, is compressed
and, optionally, dried, prior to feeding to the process of the
invention. This handling/processing is considerably safer and more
straightforward with air as opposed to oxygen enriched air or,
particularly, oxygen.
[0013] In a preferred embodiment, the air is supplied from the
atmosphere and is dried prior to entering either reactor. The air
may be dried by any drying method known in the art, but is
preferably compressed and then fed into a drying system comprising
a desiccant. Suitable desiccants include silica gel, which can dry
the air to a dew point of less than about -40.degree. C. In one
aspect there are two or more desiccant chambers so that one can be
regenerated whilst the other is drying the air.
Alternatively/additionally, the air can be cooled to condense the
water.
[0014] Typically, the amount of air co-fed to the process of the
invention is from about 0.1 to about 500 mol %, based on the amount
organics fed and/or present in the reactor(s). By organics we mean
the carbon-based compounds present in the process of the invention,
particularly 243db, 1233xf and 245cb. In one aspect, the amount of
air co-fed to the process of the invention in mol %, as described
herein, is based on the amount (i) 243db, (ii) 1233xf, or (iii) the
combined amount of 243db and 1233xf. In a preferred aspect, for
example wherein the process of the invention is carried out in
first and second reactors and the air is co-fed to the second
reactor only, the amount of air (mol %) is based on the amount of
1233xf fed to the second reactor.
[0015] Preferably, the amount of air co-fed to the process of the
invention is from about 1 to about 200 mol %, from about 2 to about
100 mol %, from about 5 to about 100 mol % or from about 10 to
about 100 mol %, based on the amount of organics. The preferred
amounts of air co-fed to the process of the invention are believed
to be limited as follows. If too little air is used, inadequate
prevention and/or retardation of coking of the catalyst or
catalysts is achieved. If too much air is used, selectivity for the
desired products is adversely affected and/or the large quantities
of air become more difficult, and therefore expensive, to handle.
It is particularly advantageous that the amount of air co-fed to
the process is from about 15 to about 95 mol %, preferably from
about 20 to about 90 mol %, such as from about 25 to about 85 mol
%, based on the amount of organics. These ranges are currently
thought to be optimal from the perspective of a combination of ease
of handling (e.g. the volume of air to be handled and its effect on
the process design) and ability to prevent and/or retard catalyst
coking without deleteriously affecting the process chemistry.
[0016] When the catalytic dehydrochlorination of 243db and the
catalytic fluorination of 1233xf are carried out in separate first
and second reactors, respectively, air may be co-fed to the first
reactor and/or the second reactor. In one embodiment, air is co-fed
to the first reactor and the second reactor, more preferably to the
second reactor only.
[0017] Thus, the invention provides a process for preparing 245cb,
the process comprising gas phase catalytic dehydrochlorination in a
first reactor of a composition comprising 243db to produce an
intermediate composition comprising 1233xf, HF, HCl; and gas phase
catalytic fluorination with HF in a second reactor of the
intermediate composition to produce a reactor product composition
comprising 245cb, HF, HCl and air; wherein the process is carried
out with a co-feed of air to the second reactor.
[0018] Whether air is co-fed to the first reactor and the second
reactor, or to the second reactor only, the amount of air co fed to
the reactor(s) is broadly in accordance with the ranges as defined
hereinbefore.
[0019] However, when air is co-fed to both the first and second
reactors, the amount of air co-fed to the first reactor preferably
is less than the amount, on a molar basis, of air co-fed to the
second reactor. This is because 1233xf typically is fouling to the
gas phase fluorination catalyst in the second reactor and higher
concentrations of air are thought to be needed to maintain catalyst
stability and activity (e.g. by preventing and/or retarding
catalyst coking) in the second reactor compared to the first
reactor. Additionally, more forcing conditions may be employed in
the second reactor compared to the first reactor in order to
achieve the desired levels of 1233xf fluorination conversion and
selectivity to 245cb. Higher concentrations of air in the second
reactor compared to the first reactor can help maintain catalyst
stability and activity under such forcing conditions.
[0020] Typically, the amount of air co-fed to the first reactor is
less than half the amount co-fed to the second reactor, preferably
less than a quarter of the amount of air co-fed to the second
reactor, such as less than a tenth of the amount of air co-fed to
the second reactor. By way of example, when air is co-fed to both
the first and second reactors, the amount of air co-fed to the
first reactor typically is from about 0.1 to about 100 mol %,
preferably from about 0.2 to about 50 mol %, such as from about 0.3
to about 20 mol %, for example from about 0.4 to about 10 mol %,
based on the amount of organics (e.g. based on 243db); whereas the
amount of air co-fed to the second reactor typically is from about
1 to about 200 mol %, preferably from about 5 to about 100 mol %,
such as from about 10 to about 90 mol %, for example from about 15
to about 85 mol %, based on the amount of organics (e.g. based on
1233xf).
[0021] In a preferred embodiment when the catalytic
dehydrochlorination of 243db and the catalytic fluorination of
1233xf are carried out in separate first and second reactors, the
intermediate composition exiting the first reactor is fed directly
to the second reactor. This has the advantage of process economy.
For example, when air is co-fed to the first reactor, the
intermediate composition contains air. It is preferable, when air
is co-fed to the first reactor, for air also to be fed to the
second reactor. This can be achieved simply by feeding the
intermediate composition exiting the first reactor directly to the
second reactor without an intermediate purification step (e.g. to
remove air and/or Hel). Is has been found that the presence of HCl
does not significantly disadvantage the fluorination of 1233xf. The
unexpected benefit of this is the intermediate composition exiting
the first reactor can be fed directly to the second reactor without
removing HCl, which removal requires energy to cool the
composition, remove HCl and re-heat the composition. Of course,
even when feeding the intermediate composition exiting the first
reactor directly to the second reactor without an intermediate
purification step, it may be desirable to heat or cool the
intermediate composition, for example if the fluorination reaction
in the second reactor is being carried out at a higher temperature
than the dehydrochlorination reaction in the first reactor. In the
embodiment wherein the intermediate composition is fed directly to
the second reactor, it is preferable to have an additional co-feed
of air into the second reactor because, as explained above, higher
concentrations of air in the second reactor compared to the first
reactor can help prevent and/or retard catalyst coking.
[0022] The catalyst used in the catalytic dehydrochlorination step
may be any suitable catalyst that is effective to dehydrochlorinate
243db. Preferred catalysts are bulk form or supported catalysts
comprising activated carbon, a zero-valent metal, a metal oxide, a
metal oxyhalide, a metal halide, or mixtures of the foregoing.
[0023] For the avoidance of doubt, by bulk form or supported
catalysts, catalysts comprising activated carbon, a zero-valent
metal, a metal oxide, a metal oxyhalide, a metal halide, or
mixtures of the foregoing, we include catalysts that are
essentially only bulk form or supported catalysts, catalysts
comprising activated carbon, a zero-valent metal, a metal oxide, a
metal oxyhalide, a metal halide, or mixtures thereof, and such
catalysts that are modified, for example, by the addition of one or
more promoters or excipients. Suitable promoters include metals
(e.g. transition metals) and/or compounds thereof, and suitable
excipients include binders and/or lubricants.
[0024] By "activated carbon", we include any carbon with a
relatively high surface area such as from about 50 to about 3000
m.sup.2 or from about 100 to about 2000 m.sup.2 (e.g. from about
200 to about 1500 m.sup.2 or about 300 to about 1000 m.sup.2). The
activated carbon may be derived from any carbonaceous material,
such as coal (e.g. charcoal), nutshells (e.g. coconut) and wood.
Any form of activated carbon may be used, such as powdered,
granulated and pelleted activated carbon. Activated carbon which
has been modified (e.g. impregnated) by the addition of Cr, Mn, Au,
Fe, Sn, Ta, Ti, Sb, AI, Co, Ni, Mo, Ru, Rh, Pd and/or Pt and/or a
compound (e.g. a halide) of one or more of these metals may be
used.
[0025] Suitable catalysts comprising a zero-valent metal including
supported (e.g. by carbon) transition metals such as Pd, Fe, Ni and
Co.
[0026] Suitable metals for the catalysts comprising a metal oxide,
a metal oxyhalide or a metal halide include transition metals,
alkaline earth metals (e.g. Mg) and main group metals such as AI,
Sn or Sb.
[0027] Alumina which has been modified by the addition of Cr, Cu,
Zn, Mn, Au, Fe, Sn, Ta, Ti, Sb, In, Co, Ni, Mo, Ru, Rh, Pd and/or
Pt and/or a compound (e.g. a halide) of one or more of these metals
may be used.
[0028] A further group of preferred catalysts are supported (e.g.
on carbon) lewis acid metal halides, including TaX.sub.5,
SbX.sub.5, SnX.sub.4, TiX.sub.4, FeCl.sub.3, NbX.sub.5, VX.sub.5,
AIX.sub.3 (wherein X=F or CI). An oxide of a transition metal that
has been modified by the addition of Cr, Mn, Au, Fe, Sn, Ta, Ti,
Sb, In, AI, Co, Ni, Nb, Mo, Ru, Rh, Pd and/or Pt and/or a compound
(e.g. a halide) of one or more of these metals may be used.
[0029] A preferred oxide of a transition metal is an oxide of Cr,
Ti, V, Zr, or Fe. For example, chromia (Cr.sub.20.sub.3) alone or
chromia that has been modified by the addition of Zn, Mn, Mo, Nb,
Zr, In, Ni, AI and/or Mg and/or a compound of one or more of these
metals may be used. Catalysts based on chromia currently are
particularly preferred. A preferred chromia-based catalyst is a
zinc/chromia catalyst.
[0030] By the term "zinc/chromia catalyst" we mean any catalyst
comprising chromium or a compound of chromium and zinc or a
compound of zinc. Such catalysts are known in the art, see for
example EP-A-0502605, EP-A-0773061, EP-A-0957074, WO 98/10862, WO
2010/116150, which documents are incorporated herein by
reference.
[0031] Typically, the chromium or compound of chromium present in
the zinc/chromia catalysts of the invention is an oxide,
oxyfluoride or fluoride (preferably an oxide or oxyfluoride) of
chromium.
[0032] The total amount of the zinc or a compound of zinc present
in the zinc/chromia catalysts of the invention is typically from
about 0.01% to about 25%, preferably 0.1% to about 25%,
conveniently 0.01% to 6% zinc, and in some embodiments preferably
0.5% by weight to about 25% by weight of the catalyst, preferably
from about 1 to 10% by weight of the catalyst, more preferably from
about 2 to 8% by weight of the catalyst, for example about 4 to 6%
by weight of the catalyst. In other embodiments, the catalyst
conveniently comprises 0.01% to 1%, more preferably 0.05% to 0.5%
zinc. It is to be understood that the amount of zinc or a compound
of zinc quoted herein refers to the amount of elemental zinc,
whether present as elemental zinc or as a compound of zinc.
[0033] The zinc/chromia catalysts used in the present invention may
be amorphous. By this we mean that the catalyst does not
demonstrate substantial crystalline characteristics when analysed
by, for example, X-ray diffraction. Alternatively, the catalysts
may be partially crystalline. By this we mean that from 0.1 to 50%
by weight of the catalyst is in the form of one or more crystalline
compounds of chromium and/or one or more crystalline compounds of
zinc. If a partially crystalline catalyst is used, it preferably
contains from 0.2 to 25% by weight, more preferably from 0.3 to 10%
by weight, still more preferably from 0.4 to 5% by weight of the
catalyst in the form of one or more crystalline compounds of
chromium and/or one or more crystalline compounds of zinc.
[0034] The percentage of crystalline material in the catalysts of
the invention can be determined by any suitable method known in the
art. Suitable methods include X-ray diffraction (XRD) techniques.
When X-ray diffraction is used the amount of crystalline material
such as the amount of crystalline chromium oxide can be determined
with reference to a known amount of graphite present in the
catalyst (e.g. the graphite used in producing catalyst pellets) or
more preferably by comparison of the intensity of the XRD patterns
of the sample materials with reference materials prepared from
suitable internationally recognised standards, for example NIST
(National Institute of Standards and Technology) reference
materials.
[0035] The zinc/chromia catalysts typically have a surface area of
at least 50 m.sup.2/g and preferably from 70 to 250 m.sup.2/g and
most preferably from 100 to 200 m.sup.2/g before it is subjected to
pretreatment with a fluoride containing species such as hydrogen
fluoride or a fluorinated hydrocarbon. During this pre-treatment,
which is described in more detail hereinafter, at least some of the
oxygen atoms in the catalyst are replaced by fluorine atoms.
[0036] The amorphous zinc/chromia catalysts which may be used in
the present invention can be obtained by any method known in the
art for producing amorphous chromia-based catalysts. Suitable
methods include co-precipitation from solutions of zinc and
chromium nitrates on the addition of ammonium hydroxide.
Alternatively, surface impregnation of the zinc or a compound
thereof onto an amorphous chromia catalyst can be used.
[0037] Further methods for preparing the amorphous zinc/chromia
catalysts include, for example, reduction of a chromium (VI)
compound, for example a chromate, dichromate, in particular
ammonium dichromate, to chromium (III), by zinc metal, followed by
co-precipitation and washing; or mixing as solids, a chromium (VI)
compound and a compound of zinc, for example zinc acetate or zinc
oxalate, and heating the mixture to high temperature in order to
effect reduction of the chromium (VI) compound to chromium (III)
oxide and oxidise the compound of zinc to zinc oxide.
[0038] The zinc may be introduced into and/or onto the amorphous
chromia catalyst in the form of a compound, for example a halide,
oxyhalide, oxide or hydroxide depending at least to some extent
upon the catalyst preparation technique employed. In the case where
amorphous catalyst preparation is by impregnation of a chromia,
halogenated chromia or chromium oxyhalide, the compound is
preferably a water-soluble salt, for example a halide, nitrate or
carbonate, and is employed as an aqueous solution or slurry.
Alternatively, the hydroxides of zinc and chromium may be
co-precipitated (for example by the use of a base such as sodium
hydroxide or ammonium hydroxide) and then converted to the oxides
to prepare the amorphous catalyst. Mixing and milling of an
insoluble zinc compound with the basic chromia catalyst provides a
further method of preparing the amorphous catalyst precursor. A
method for making amorphous catalyst based on chromium oxyhalide
comprises adding a compound of zinc to hydrated chromium
halide.
[0039] The amount of zinc or a compound of zinc introduced to the
amorphous catalyst precursor depends upon the preparation method
employed. It is believed that the working catalyst has a surface
containing cations of zinc located in a chromium-containing
lattice, for example chromium oxide, oxyhalide, or halide lattice.
Thus the amount of zinc or a compound of zinc required is generally
lower for catalysts made by impregnation than for catalysts made by
other methods such as co-precipitation, which also contain the zinc
or a compound of zinc in non-surface locations.
[0040] The catalysts described herein (e.g. the chromia-based
catalysts such as zinc/chromia catalysts) are typically stabilised
by heat treatment before use such that they are stable under the
environmental conditions that they are exposed to in use. This
stabilisation is often a two-stage process. In the first stage, the
catalyst is calcined by heat treatment in nitrogen or a
nitrogen/air environment. The catalyst is then typically stabilised
to hydrogen fluoride by heat treatment in hydrogen fluoride. This
stage is often termed "prefluorination".
[0041] By careful control of the conditions under which these two
heat treatment stages are conducted, crystallinity can be induced
into the catalyst to a controlled degree.
[0042] In use, the catalysts described herein (e.g. the
chromia-based catalysts such as the zinc/chromia catalysts) may be
regenerated or reactivated periodically by heating in air at a
temperature of from about 300.degree. C. to about 500.degree. C.
Air may be used as a mixture with an inert gas such as nitrogen or
with hydrogen fluoride, which emerges hot from the catalyst
treatment process and may be used directly in any fluorination
processes employing the reactivated catalyst.
[0043] The vapour phase catalytic dehydrochlorination may be
carried out at a temperature of from about 200 to about 450.degree.
C. and at atmospheric, sub- or super-atmospheric pressure,
preferably from about 0.1 to about 30 bara. Preferably, the
catalytic dehydrochlorination is conducted at a temperature of from
about 250 to about 400.degree. C., such as from about 280 to about
380.degree. C. or from about 300 to about 350.degree. C.
[0044] The vapour phase catalytic dehydrochlorination preferably is
carried out at a pressure of from about 0.5 to about 25 bara or
about 1 to about 20 bara, such as from about 2 to about 18 bara
(e.g. about 5 to about 20 bara or about 8 to about 18 bara or about
10 to about 15 bara).
[0045] HF is required for the fluorination of 1233xf in the process
of the invention. The molar ratio of HF: 1233xf in the catalytic
fluorination step is typically from about 1:1 to about 45:1, such
as from about 1:1 to about 30:1, preferably from about 1.5:1 to
about 30:1, such as from about 2:1 to about 20:1 or from about 3:1
to about 15:1. The inventors have unexpectedly found that these
ranges strike a balance between the desirability to prevent and/or
retard catalyst coking and residence time. If too little HF is
used, coking increases. If too much HF is used, the residence time
for a given reactor volume becomes shorter than desired.
[0046] If the catalytic dehydrochlorination and fluorination
reactions are carried out in the same reactor, then both reactions
are carried out in the presence of HF. When first and second
reactors are used for the catalytic dehydrochlorination and
fluorination reactions, then there need not be any HF present in
the first reactor for the catalytic dehydrochlorination reaction.
However, in some embodiments it is thought preferable to have HF
present for the catalytic dehydrochlorination. Without being bound
by theory, this is believed to prevent and/or retard catalyst
coking.
[0047] If HF is present in the catalytic dehydrochlorination step,
the molar ratio of HF:243db can fall within the ranges defined
above for the molar ratio of HF:1233xf in the catalytic
fluorination of 1233xf. In one aspect, however, less HF is used in
the catalytic dehydrochlorination step compared to the catalytic
fluorination step. Thus, the composition comprising 243db can
additionally contain HF, typically in a molar ratio of HF:243db of
from about 0.5:1 to about 40:1, such as from about 0.5:1 to about
20:1, preferably from about 1:1 to about 15:1, such as from about
1.5:1 to about 10:1 or from about 2:1 to about 8:1.
[0048] The contact time for the composition comprising 243db and HF
with the catalyst in the catalytic dehydrochlorination step
typically is from about 0.5 to about 200 seconds, such as from
about 1 to about 150 seconds. Preferably, the contact time is from
about 1 to about 100 seconds, such as from about 2 to about 80
seconds or from about 8 to about 60 seconds.
[0049] Turning now to the gas phase catalytic fluorination of the
intermediate composition of the process of the invention, the HF in
the intermediate composition typically is used to fluorinate 1233xf
to 245cb. Preferably, the HF in the intermediate composition is the
sole fluorinating agent for conversion of 1233xf to 245cb, although
additional HF can be added to the process of the invention to
facilitate this, particularly if a second reactor is used for the
catalytic fluorination of 1233xf.
[0050] The catalyst used in the catalytic fluorination step may be
any suitable catalyst that is effective to fluorinate 1233xf to
245cb. Preferred catalysts are bulk form or supported catalysts
comprising activated carbon, a zero-valent metal, a metal oxide, a
metal oxyhalide, a metal halide, or mixtures of the foregoing as
described above in relation to the catalyst for the catalytic
dehydrochlorination step.
[0051] Preferred catalysts for catalytic fluorination of 1233xf to
245cb are those which comprise chromia, alone or chromia that has
been modified by the addition of Zn, Mn, Mo, Nb, Zr, In, Ni, AI
and/or Mg and/or a compound of one or more of these metals. A
preferred chromia-based catalyst for use in the catalytic
fluorination of 1233xf to 245cb is a zinc/chromia catalyst. The
same catalyst (e.g. a chromia-based catalyst) may be used for the
catalytic dehydrochlorination and fluorination steps.
[0052] The vapour phase catalytic fluorination step may be carried
out at a temperature of from about 200 to about 450.degree. C. and
at atmospheric, sub- or super-atmospheric pressure, preferably from
about 0.1 to about 30 bara. Preferably, the vapour phase catalytic
fluorination is conducted at a temperature of from about 250 to
about 420.degree. C., such as from about 280 to about 400.degree.
C. or from about 300 to about 380.degree. C. (e.g. from about 330
to about 380.degree. C.)
[0053] The vapour phase catalytic fluorination preferably is
carried out at a pressure of from about 0.5 to about 25 bara or
about 1 to about 20 bara, such as from about 2 to about 20 bara
(e.g. about 5 to about 20 bara or from about 10 to about 15
bara).
[0054] The contact time for the for the composition comprising
1233xf, HCl and HF with the catalyst in the catalytic fluorination
step typically is from about 0.5 to about 200 seconds, such as from
about 1 to about 150 seconds. Preferably, the contact time is from
about 1 to about 100 seconds, such as from about 2 to about 80
seconds or from about 5 to about 50 seconds.
[0055] 245cb is a useful starting material for the manufacture of
1234yf. Accordingly, the process of the invention further comprises
feeding 245cb into a dehydrofluorination reactor to produce a
dehydrofluorination product comprising 2,3,3,3-tetrafluoropropene
(1234yf) and HF.
[0056] The dehydrofluorination of 245cb may be carried out in the
vapour and/or liquid phase and typically is carried out at a
temperature of from about -70 to about 1000.degree. C. (e.g. 0 to
450.degree. C.). The dehydrofluorination may be carried out at
atmospheric sub- or super atmospheric pressure, preferably from
about 0.1 to about 30 bara.
[0057] The dehydrofluorination may be induced thermally, may be
base-mediated and/or may be catalysed by any suitable catalyst.
Suitable catalysts include metal and carbon based catalysts such as
those comprising activated carbon, main group (e.g. alumina-based
catalysts) and transition metals, such as chromia-based catalysts
(e.g. zinc/chromia), lewis acid metal halides or zero-valent metal
catalysts. One preferred method of effecting the
dehydrofluorination of the compound of 245cb to produce 1234yf is
by contacting with a metal-based catalyst, such as a chromia-based
(e.g. zinc/chromia) catalyst.
[0058] Preferably, the 245cb is catalytically dehydrofluorinated to
1234yf in the gas phase.
[0059] 243db is commercially available (e.g. from Apollo Scientific
Ltd, UK). Alternatively, 243db may also be prepared via a synthetic
route starting from the cheap feedstocks carbon tetrachloride
(CCI4) and ethylene (see the reaction scheme set out below). These
two starting materials may be telomerised to produce
1,1,1,3-tetrachloropropane (see, for example, J. Am. Chem. Soc.
Vol. 70, p2529, 1948, which is incorporated herein by reference)
(also known as HCC-250fb, or simply 250fb).
[0060] 250fb may then be fluorinated to produce
3,3,3-trifluoropropene (1243zf) and/or 1,1,1-
trifluoro-3-chloropropane (253fb) (e.g. using HF, optionally in the
presence of a chromiacontaining catalyst, preferably a zinc/chromia
catalyst as described herein). Dehydrohalogenation of
1,1,1-trifluoro-3-chloropropane (e.g. using NaOH or KOH or in the
vapour phase) produces 3,3,3-trifluoropropene (1243zf).
[0061] 1243zf may then be readily halogenated, such as chlorinated
(e.g. with chlorine) to produce 1,1,1-trifluoro-2,3-dichloropropane
(243db). This reaction scheme is summarized below.
##STR00001##
[0062] The preparation of 243db outlined above is described in more
detail in WO 2010/116150 and WO 2009/125199, which are incorporated
herein by reference.
[0063] Embodiments of the present invention will now be described
with reference to the following non-limiting examples and
drawings:
DESCRIPTION OF THE DRAWINGS
[0064] FIG. 1 shows a schematic process flow sheet in accordance
with the invention;
[0065] FIG. 2 shows the results of a coking study in which
conversion is plotted over time for the fluorination of 1233xf in
accordance with the invention.
DETAILED DESCRIPTION
[0066] FIG. 1 illustrates a process design in accordance with the
invention. A composition (1) comprising 243db and HF is introduced
into a first reactor (A) in which gas phase catalytic
dehydrochlorination occurs to produce an intermediate composition
(2) comprising 1233xf, HF and HCl. The intermediate composition may
further contain unreacted 243db and, in certain embodiments,
by-products such as 245cb and 1234yf.
[0067] The intermediate composition (2) is fed directly to a second
reactor (B), as is a co-feed (3) of air, and gas phase catalytic
fluorination of the intermediate composition (2) occurs in the
second reactor (B) to produce a reactor product composition (4)
comprising 245cb, HF, HCl and air. The reactor product composition
may further contain unreacted 1233xf and, in certain embodiments,
unreacted 243db and by-products such as 1234yf.
[0068] In a preferred embodiment, the reactor product composition
(4) is separated at separation step (C) into a stream (5)
comprising HCl and air and a stream (6) comprising 245cb and HF. An
advantage of the use of the co-feed of air in the process of the
invention is that it can be readily separated from the reactor
product composition together with HCl. Preferably, this is achieved
by distillation, with the stream (5) comprising HCl and air being
taken off the top of a distillation column (e) and the stream (6)
comprising 245cb and HF being taken off the bottom of the
distillation column (e). The stream (6) typically contains any
other components present, such as unreacted 243db, 1233xf and/or
1234yf.
[0069] In the embodiment illustrated by FIG. 1, the stream (6)
comprising 245cb and HF is separated at separation step (D) into a
245cb-rich stream (7) and a HF-rich stream (8). Preferably, this is
achieved by distillation, with the 245cb-rich stream (7) being
taken off the top of a distillation column (D) and the HF-rich
stream (8) being taken off the bottom of the distillation column
(D). The 245cb-rich stream (7) typically also contains any
relatively light organic components present, such as 1234yf. The
HF-rich stream (8) typically also contains any relatively heavy
organic components present, such as 1233xf.
[0070] Preferably the 245cb-rich stream (7) is subjected to a
scrubbing step (E) in which any residual HF (and/or indeed any
residual HCl) is substantially removed from the 245cb-rich stream
to produce a 245cb-rich stream (11) substantially free from HF
(and/or substantially free from HCl). Typically, this step (E)
involves contacting the 245cb-rich stream (7) with water and/or or
with a source of aqueous acid and/or or with a source of aqueous
alkali, generally represented in FIG. 1 as stream (9), to generate
the 245cb-rich stream (11) substantially free from HF and one or
more spent scrubbing streams (10). By substantially free from HF,
we include the meaning of less than 100 ppm, preferably less than
50 ppm, 40 ppm, 30 ppm, 20 ppm, 10 ppm, 5 ppm 4, ppm, 3 ppm, 3 ppm
or less than 1 ppm. In a preferred embodiment, the 245cb-rich
stream (11) is subjected to a separation step (F) in which the
245cb is further separated from any further organic components
present (e.g. fluorocarbons such as 1234yf) to produce a
substantially pure 245cb product (13). Preferably, this separation
step (F) comprises one or more distillation steps. By substantially
pure 245cb product (13), we include the meaning of greater than
95%, 98%, 99% pure, preferably greater than 99.5%, 99.8% or 99.9%
pure, on a molar basis.
[0071] In a preferred embodiment, the HF in the HF-rich stream (8)
is recycled to the catalytic dehydrochlorination of the composition
comprising 243db and HF. As shown in FIG. 1, preferably, the
HF-rich stream is subjected to a separation step (G) in which the
HF-rich stream (8) is separated into an HF stream (14) and an
organic stream (15). The HF stream is recycled to the composition
(1) comprising 243db and HF which enters the first reactor (A) in
which gas phase catalytic dehydrochlorination occurs. In a
preferred aspect, the separation step (G) comprises a phase
separator.
EXAMPLE 1
[0072] A series of catalysts (see Table 1 below) were screened for
243db dehydrochlorination. The test catalysts were ground to
0.5-1.4 mm and 2 mL was charged to an Inconel 625 reactor (0.5''
OD.times.32 cm). The catalysts were pre-dried at 200.degree. C. for
at least 2 hours under a flow of N2 (60 ml/min). All the catalysts
shown, except activated carbon, were pre-fluorinated as follows. HF
at 30 ml/min was passed over the catalyst along with 60 ml/min
nitrogen at 300.degree. C. for one hour. The nitrogen was directed
to the reactor exit leaving neat HF passing over the catalyst. The
temperature was slowly ramped to 360.degree. C. and held for hours
before reducing to 250.degree. C. All the experiments were run at
atmospheric pressure and at the temperatures indicated. The 243db
flow was 2ml/min with activated carbon catalyst and ranged from 0.5
to about 1 ml/min for the remaining catalysts. All experiments were
conducted with an HF flow in excess of the 243db flow, except for
the activated carbon catalyst runs, in which no HF was used.
Reactor off-gas was sampled scrubbing through deionised water and
analysed by gas chromatography. The 243db conversion and
selectivity to 1233xf are shown in Table 1.
TABLE-US-00001 TABLE 1 Experimental results for 243db
dehydrochlorination Temperature 243db 1233xf Catalyst .degree. C.
Conversion % selectivity % 2% Zn/Chrome 250 98.22 74.79 300 100.00
42.28 350 100.00 68.74 4% Zn/Chrome 250 99.18 81.80 300 98.47 77.42
350 100.00 70.80 6% Zn/Chrome 250 94.94 65.55 300 100.00 33.81 350
100.00 71.30 8% Zn/Chrome 250 89.06 60.20 300 100.00 82.93 350
100.00 72.72 Chrome 250 56.69 46.72 300 98.93 83.35 350 100.00
72.31 5% In/Chrome 250 98.64 81.36 300 100.00 68.23 350 100.00
52.65 6% Zn/Chrome 250 96.51 85.17 300 100.00 69.18 350 100.00
72.19 Zn/Chrome 250 95.64 82.78 300 100.00 71.90 350 100.00 70.47
Zn/Chrome 250 91.76 81.19 300 100.00 77.74 350 100.00 70.90 Chrome
250 93.60 80.31 300 100.00 48.69 350 100.00 38.38 Mo/Chrome 250
93.24 80.25 300 100.00 54.87 350 100.00 26.63 Ni/Chrome 250 100.00
84.05 300 100.00 39.93 350 100.00 24.93 Nb/Chrome 250 98.94 86.54
300 100.00 56.26 350 100.00 34.94 Alumina 250 29.10 38.87 300 73.37
69.75 350 98.37 90.94 0.5% Pt/Alumina 250 44.53 48.63 300 87.56
82.09 350 100.00 96.02 Fe/Alumina 250 22.45 39.46 300 52.57 71.90
350 80.99 84.47 20% Cr/Alumina 250 43.85 44.17 300 97.71 79.58 350
98.56 77.88 50% Cr/Alumina 250 45.73 42.68 300 100.00 79.86 350
100.00 72.95 Zn/Cu/Alumina 250 40.73 47.65 300 74.85 71.21 350
65.17 64.86 0.5% Pd/Carbon 250 100.00 99.19 300 100.00 98.38 350
100.00 85.78 activated carbon 175 46.06 99.13 200 98.46 97.32 300
99.07 97.02
All of the catalysts tested were found to be effective at
converting 243db to 1233xf, particularly activated carbon.
EXAMPLE 2
[0073] 6.07 g of a Indium-doped chromia catalyst was dried over 72
hours under nitrogen (80 ml/min) at 250.degree. C. and 3 barg. This
was followed by two-stage pre-fluorination of the catalyst. In
stage 1, the catalyst was exposed to nitrogen (80 ml/min) and HF (4
ml/min) at 250.degree. C. and 3 barg up until 4 hours from HF
breakthrough, at which time the temperature was increased at
25.degree. C./min to 300.degree. C. and held for 16 hours. In stage
2, nitrogen flow was reduced stepwise until it was switched off,
and the temperature was increased at 25.degree. C./min to
380.degree. C. and held for 10 hours. The HF flow was stopped and
replaced with nitrogen (40 ml/min) and the temperature reduced to
250.degree. C. ready for use.
[0074] 1233xf was co-fed with HF over the catalyst without an air
co-feed (cycle 1) and with an air co-feed (cycle 2) for about 100
hours at 350.degree. C. and 15 barg. Reactor off-gas was analysed
by GC. Monitored catalyst regeneration was used to measure the
average coke levels in the catalyst after use. The results are
shown in Table 2 below and illustrated in FIG. 2.
TABLE-US-00002 TABLE 2 Experimental conditions and results for two
ageing runs Conversion Cycle Coke Loss time levels Target flows
(ml/min) Loss rate Cycle (hrs) (%) HF 1233xf Air (%) Hours (5/hr) 1
100.5 5.6 50 5 -- 64.8 74 0.88 2 110 0.35 45 5 5 16.8 106 0.16
[0075] Both cycles were conducted at the same temperature and
pressure, but the HF flow was reduced in cycle 2 to maintain
contact time. Contact times for cycles 1 and 2 were 57 seconds and
65 seconds, respectively. As a result of the reduced HF flow on
cycle 2 and the lower than target 1233xf flows, which were hard to
control and lower than target, the HF:1233xf ratio differed
slightly on the two cycles (average 20:1 for cycle 1 and 15:1 for
cycle 2). The average 1233xf flow for cycle 1 was 3.2 ml/min and
3.5 ml/min for cycle 2.
[0076] Without air the majority of catalyst activity was lost after
about 80 hours. The introduction of air significantly reduced the
rate of catalyst deactivation (the activity loss after 100 hours
was comparable with just 20 hours without air). Based on this
conversion loss rate, cycle 2 would be expected to take 410 hours
to reach the same conversion loss as cycle 1. The reduced rate of
catalyst deactivation with air co-feed is in accordance with the
catalyst coke levels measured.
[0077] The reaction selectively was also affected, total impurity
levels approximately doubled with the co-feed of air compared to
without air. Co-feeding air seemed to have a little impact on the
245cb:1234yf ratio though.
[0078] The concentration of air present was higher than desired
because the 1233xf flow rate was, on average, lower than the
targeted 5 ml/min. This was thought to be at least partially
responsible for the decreased selectivity. For this reason, and
based on the coke produced in cycle 1, a lower air concentration is
believed to be desirable to achieve comparable reduced rates of
catalyst deactivation without reducing conversion and 245cb
selectivity. It was estimated that lower air flows, for example
from about 0.5 ml/min to about 4.5 ml/min, preferably from about 1
to about 4 ml/min, such as from about 1.5 to 3.5 ml/min (all based
on an actual 1233xf flow of 5 ml/min) would realise the surprising
balance of reduced rates of catalyst deactivation combined with
conversion selectivity to the desired 245cb product.
[0079] The invention is defined by the following claims.
* * * * *